EA009365B1 - Method of operating a riser reactor - Google Patents

Abstract

A method for maximizing the contact between solid catalyst particles and a feedstock while minimizing undesirable side reactions is provided. A plurality of feedstock jets are contacted with solid catalyst particles in a reactor having a nonhomogenous contact zone and a unified flow contact zone.

Description

The present invention relates to a method for controlling reactivity in a reactor for carrying out gas-solid phase reactions (reactions between gases and solid particles). More specifically, the present invention relates to a method for converting oxygenates to olefins, which optimizes the yield of reactions converting oxygenates (oxygen-containing hydrocarbons) to olefins.

In industrial reactors, light olefins, in particular ethylene and propylene, are usually obtained as the end products from the conversion of hydrocarbons. When producing light olefins under industrial conditions, it is desirable that the conversion of the starting oxygenates to olefins occurs almost completely. In fact, however, for a number of reasons, an increase in the reaction rate of the conversion of oxygenates to light olefins usually leads to a decrease in its selectivity and, conversely, an increase in selectivity usually leads to a decrease in the reaction rate.

I8 6166282 B1 describes a method for converting oxygenates to olefins in a reactor with two reaction zones. In this reactor, the feed is fed through one pipe into the lower zone filled with the dense phase, in which the partial conversion of hydrocarbons to olefins takes place. The reaction in the first zone proceeds in a fluidized bed of catalyst. The feed to the reactor rises and enters the transition zone located above the catalyst bed, in which the substantially complete termination of the ongoing conversion reaction occurs. In the first dense zone filled with a fluidized bed of catalyst, conversion occurs at a reduced velocity of less than 1 m / s, and in the transition zone at a reduced velocity of 1 to 4 m / s.

Currently, there is a need for additional methods for increasing the efficiency of reactors designed to convert oxygenates to olefins, and other fluidized bed reactors. So, in particular, it is necessary to ensure the possibility of supplying feedstock through the side walls of the reactor. The creation of such reactors would allow, at the maximum efficiency of the conversion of oxygenates to olefins, to simultaneously increase the selectivity of the reaction and increase the yield of ethylene and propylene.

The present invention provides a method for converting starting oxygenates to olefins (olefin products). When implementing this method, the stream of catalyst forming olefins is directed through the channel for introducing the catalyst into contact with the oxidate (channel for introducing into the contact) and separate streams of oxygenate are fed into this stream at an angle to its direction, as a result of which the conversion of oxygenate to olefins begins. In this contacting channel, the individual jets of oxygenate are mixed with each other below their place in the catalyst stream relative to the direction of this stream within the specified contact channel. Next, the resulting mixture is sent to the connecting channel, which has a cross-sectional area less than the cross-sectional area of the contact channel and in which the conversion of oxygenate to olefins continues.

The olefin-forming catalyst and oxygenate from the specified contact channel may enter the uniform contact zone connected to it within the channel. The olefin-forming catalyst and oxygenate continuously passing through this uniform contact zone then pass through said connecting channel.

In a preferred embodiment, the olefin-forming catalyst stream is introduced into the contacting channel from the inlet channel, the contacting channel having a section that expands in cross section.

The expanding channel may include a non-uniform contact zone and a uniform contact zone. In an inhomogeneous contact zone, the olefin-forming catalyst interacts with the oxygenate, and then the olefin-forming catalyst and oxygenate enter the uniform contact zone.

In particular embodiments, individual oxygenate jets are injected into the contact channel at a speed of 20 to 40 m / s.

The concentration of the olefin-forming catalyst in the connecting channel is preferably lower than in the contacting channel.

The solids content in the olefin-forming catalyst in the contact channel can be from 15 to 65 vol.%, And in the connecting channel from 3 to 15 vol.%, In particular from 1 to 5 vol.%.

The obtained olefins are further separated from the catalyst forming them, and at least one obtained olefin contained in the olefins separated from the catalyst is treated with an olefin forming catalyst to obtain a polyolefin.

In the contact channel, less than 100% of the total oxygenate fed to it is preferably converted to olefins.

The methods proposed in these embodiments can produce not only olefins, but also polyolefins. The mixture of the olefin product and the catalyst removed from the reactor can then be separated into the olefin product and the catalyst. At least one olefin from this isolated olefin product can then be contacted with an olefin-forming catalyst to produce a polyolefin.

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Below the invention is described in more detail with reference to the accompanying drawings, which show:

in FIG. 1 is a schematic view of the bottom of a reactor that can be used to produce olefins by the method of the invention, FIG. 2-5 are graphs containing information about the structure of the stream in the reactor for carrying out gas-solid phase reactions upon receipt of olefins by the method of the invention.

Preferred Embodiments

I. Maximizing the efficiency of contact of oxygenate with the catalyst and minimizing adverse reactions.

The present invention provides a method for maximizing contact efficiency between a feed stream and a stream of catalyst particles while minimizing undesired side reactions. The method according to the invention allows, in particular, to maximize the contact efficiency between oxygenate compounds and catalyst particles, which converts oxygenate into olefin products consisting essentially of ethylene and propylene.

In one embodiment, particles of the olefin-forming catalyst are fed into the contacting channel through a first or inlet channel. In a preferred embodiment, the contact channel has a larger cross section than the first (input) channel.

In another embodiment, the olefin-forming catalyst particles form a fluidized catalyst bed in the contact channel. One or more streams of the starting oxygenate are fed into the stream of catalyst particles passing through the contact channel through the side wall of the contact channel. The starting oxygenate is preferably injected into the contacting channel in the form of several jets at an angle to the direction of the catalyst passing through the contacting channel. Oxygenate jets passing through the fluidized bed of the catalyst interact with the catalyst and remove some of its particles from the catalyst bed. In the process of interaction of oxygenate with catalyst particles in the reactor, a conversion reaction begins.

In the fluidized bed of the catalyst, oxygenate jets injected into the reactor meet and mix with each other. Oxygenate jets meet each other at the boundary of the fluidized bed, and the flow pattern gradually changes from a pronounced jet and inhomogeneous to a homogeneous cork mode of a two-phase gas flow with catalyst particles suspended in it. A change in the nature of the flow and the structure of the flow occurs either without formation of an intermediate zone with a turbulent flow between the zone of an inhomogeneous jet flow and a zone with a uniform flow structure, or with the formation of a small turbulent zone. The transition from a jet flow to a plug flow regime of a two-phase flow with a minimum zone of turbulent flow maximizes the intensity of the interaction between oxygenate and solid particles of the catalyst. As a result of this, at a certain stream length of the catalyst and oxygenate flowing through the reactor, the oxygenate conversion increases, and the intensity of side reactions is kept to a minimum.

A change in the flow regime from an inhomogeneous jet flow regime with oxygen particles carried by solid catalyst particles to a cork homogeneous flow regime of a two-phase flow occurs within the contact channel. The zone of change in the flow regime can be considered as a zone that divides the channel of introduction into contact into two parts. The first part is heterogeneous and in it the oxygenate jets retain their structure. The flow regime in this part of the contacting channel has a pronounced jet character and is determined by individual jets of oxygenate flowing through the fluidized bed of the catalyst. The concentration of catalyst particles in the oxygenate jets is less than in the rest of the fluidized bed. Different concentrations of oxygenate determine the heterogeneous flow structure with distinct areas with high and low concentrations of catalyst particles.

At the entrance to the contact channel, the oxygenate jets are directed at an angle to the flow direction of the catalyst flowing through it. Oxygenate jets passing through the fluidized bed of the catalyst meet each other. As you move through the fluidized bed of catalyst, the cross section of oxygenate jets gradually increases. As a result, from a gradually expanding (converging with each other) jets in the channel of introduction into contact one oxygenate stream is formed. In the portion of the channel for introducing into the contact located before the finally expanded oxygenate jets, the flow structure can be characterized as a structure with separate jets separated by gaps. This part of the contacting channel in which the individual jets of oxygenate are separated by gaps is called the separation zone of the individual jets of oxygenate. This portion of the contacting channel is filled with fluidized catalyst particles. The second part of the contact channel is uniform in flow structure. The flow regime in this part is no longer of a jet nature, and oxygenate and catalyst particles form one two-phase flow.

From the contacting channel consisting of oxygenate and catalyst particles, the stream enters the connecting channel. In a preferred embodiment, the cross section of the connecting channel is less

- 2 009365 cross-section of the channel of introduction into contact. In one embodiment, less than 100% of the starting oxygenate is converted to olefins in the contact channel. In this embodiment, gases that contain less than 100% of the initial oxygenate fall into the connecting channel from the introduction channel into contact.

After conversion, the olefins can be separated from the gas stream leaving the reactor and used, for example, to produce polyolefins. At least one of the olefins resulting from the conversion of the olefins is treated with another catalyst to obtain the desired polyolefins.

The following part of the description deals with the characteristics and structure of the streams of the initial oxygenate and catalyst particles. It must be emphasized that the reaction of the conversion of oxygenate to olefins begins immediately at the moment of contact of the initial oxygenate with catalyst particles. As the flow in the direction of exit from the reactor, the olefin content in the gas stream gradually increases. It will be apparent to those skilled in the art that after contacting the starting oxygenate with particles of the olefin-forming catalyst, the stream of gas flowing through the reactor consists of the starting oxygenate, the olefin resulting from the conversion of the starting oxygenate, and catalyst particles suspended in the gas.

II. Stream characteristics.

A. The first (input) channel.

The flow in the reactor begins when solid particles of the catalyst, in particular the catalyst forming olefins, are fed into the first or inlet channel of the reactor. Particles of fresh catalyst, particles returned to the reactor from the regenerator, or particles returned to the reactor bypassing the regenerator can be fed into the reactor. In preferred embodiments, the catalyst particles are fed into the reactor inlet channel in a fluidizing gas stream forming a fluidized catalyst bed in the reactor.

B. Channel contact.

The catalyst particles passing through the first (inlet) channel enter the contact channel. In one embodiment, the cross section of the contact channel is larger than the cross section of the inlet channel. So, for example, to increase the cross section of the contact channel, starting from the place where it is connected to the inlet, the reactor vessel can be made in the form of a confuser with inclined walls and a gradually increasing cross-sectional area.

In the contact channel, the catalyst particles interact with one or more oxygenate streams supplied to the reactor. In one embodiment, oxygenate is injected into the contact channel in the form of several jets. Oxygenate jets can be injected into the catalyst at any height in the contact channel. Inlet openings in the walls of the reactor vessel through which oxygenate is injected (injected) into it must be located at an angle to the vertical. Typically, the angle of inclination of the oxygenate jets to the vertical or direction of the catalyst flowing through the contact channel is from 15 to 75 °. The holes for supplying oxygenate jets into the contact channel are preferably made in inclined sections of the side walls of the reactor vessel.

In the contact channel, the oxygenate jets, together with the catalyst particles, form a part of the contact channel that is inhomogeneous in structure. Initially, the flow of oxygenate passing through the layer of solid particles of the catalyst is of a pronounced jet character. The flow in the non-uniform part of the contacting channel consists of separate, not containing catalyst particles, oxygenate jets surrounded by a fluidized bed of catalyst particles. The dimensions of the jets at the entrance to the contacting channel are determined by the size of the nozzles through which oxygenate is injected into the contacting channel. The oxygenate jets gradually expand in the contact channel and the catalyst particles begin to fly away from the surrounding fluidized bed. The gaps between the individual oxygenate jets in this part of the contact channel are filled with solid particles of a fluidized catalyst. As you move through the contact channel, the cross section of the jets gradually increases, and at the same time, the number of catalyst particles in the oxygenate stream increases. As a result, the concentration of catalyst particles in the oxygenate jets does not increase, and the contact between the catalyst particles and oxygenate increases.

The oxygenate injected into the contact channel at an angle to the vertical and expanding jets eventually converge with each other and form one common stream. The place in the reactor in which the individual jets of oxygenate (together with the catalyst particles) converge into a single stream determines the vertical position in the reactor of the transition section in which the structure of the stream changes from inhomogeneous to homogeneous. Typically, this transition portion is located within the contact channel. The stream of initial oxygenate and catalyst particles flowing from the second part of the contact channel into the contact moves in the reactor downward in the direction of the stream passing through the channel.

B. Connecting channel.

From the contacting channel, a stream of oxygenate and solid catalyst particles enters the connecting channel. In one embodiment, the cross-sectional area of the connecting channel is less

- 3 009365 the cross-sectional area of the contact channel. Therefore, in the connecting channel, the rate of oxygenate and catalyst particles increases. To change the cross-sectional area of the wall of the connecting channel, you can perform inclined, gradually converging to the Central axis of the channel.

The flow structure in the connecting channel is determined by the uniform structure of the flow formed in the second part of the contact channel. In preferred embodiments, the flow structure in the connecting channel is uniform and characteristic of the plug flow regime of the annular two-phase flow. In such an annular stream, solid catalyst particles are collected at the wall of the reactor, and their concentration in the central core of the stream decreases. Due to the relatively low concentration of catalyst particles in the cork biphasic stream, the reaction rate is lower than in other zones of the reactor. With a decrease in catalyst concentration, however, the selectivity of the reaction increases, and the yield of polyethylene and propylene increases compared to other olefins. The cork flow regime in the last part of the contact channel provides the necessary balance between the maximum possible conversion of oxygenate to olefins and the high selectivity of the reaction and obtaining the maximum amount of ethylene and propylene.

The connecting channel is the last part of the reaction zone through which the stream of feed oxygenate and catalyst particles passes. After that, catalyst particles are separated from the stream leaving the connecting channel, and the conversion reaction is terminated. The olefins resulting from the conversion reaction are then isolated from the gas stream purified from the catalyst particles.

G. Conditions for the reaction in the channels.

The reaction rate and reaction products of the conversion of oxygenate to olefins depend on several factors and, in particular, on the amount of coke formed on the catalyst particles and the reaction temperature. The reaction rate and the reaction products are also affected by the concentration of catalyst particles in the convertible oxygenate. For example, an increase in the density of catalyst particles accelerates the conversion of oxygenate to olefins. On the other hand, with an increase in the density of catalyst particles in the reaction products, the content of high molecular olefins and not ethylene and propylene increases. Conversely, a decrease in the density of catalyst particles decreases the reaction rate, but increases its selectivity and the amount of ethylene and propylene produced.

The necessary balance between the high reaction rate and its high selectivity can be achieved by carrying out the reaction in a mode in which the reaction can proceed under different conditions. So, for example, the reaction of the conversion of oxygenate to olefins can be carried out in a fluidized-bed reactor used for conducting gas-solid-phase reactions. In this case, the starting oxygenate is passed through a layer of catalyst particles. The relatively high concentration of catalyst particles in the fluidized bed allows an increase in the reaction rate. As oxygenate passes through the bed of catalyst particles, some of the catalyst particles are carried away by the gas stream from the fluidized bed. Due to the presence of catalyst particles, the conversion reaction continues in the gas stream passing through the catalyst bed, albeit at a lower rate due to the lower concentration of catalyst particles.

Upon receipt of olefins by the method of the invention, the flow regime in the reactor of particles of the olefin-forming catalyst and the initial oxygenate changes several times. In the first or inlet channel, the velocity of the catalyst particles is from about 0.1 to 3 ft / s, preferably from 0.5 to 1.5 ft / s. The volumetric content of the solid phase in the stream passing through the inlet channel is from about 20 to 40%, preferably from 30 to 35%. The velocity and concentration of the catalyst particles in the inlet channel is partially controlled by varying the velocity of the fluidizing gas used to form the fluidized bed of the catalyst.

The speed and concentration of catalyst particles practically do not change until the contact of the catalyst particles in the fluidized bed with oxygenate streams in the contact channel. In one embodiment, the initial oxygenate velocity at the entrance to the contact channel is from about 75 to 125 ft / s, preferably from 90 to 110 ft / s.

During the interaction of the oxygenate jets with the catalyst particles in the fluidized bed, the oxygenate velocity begins to decrease, and gradually expanding oxygenate jets carry away the catalyst particles from the fluidized bed. This process continues throughout the entire first part of the contact channel with the heterogeneous flow structure. At the upper boundary of the first contact zone, the oxygenate velocity typically decreases to about 18-30 ft / s. At the same time, the concentration of solid catalyst particles in the oxygenate stream increases to 15 vol.%. In this case, however, the concentration of solid particles in the oxygenate stream remains below their concentration in the surrounding fluidized-bed stream. In one embodiment, the concentration of solid catalyst particles in the jets of oxygenate is from about 1 to 15 vol.%. The concentration of solid particles of the catalyst between the jets remains at the level of 5% of the volume of the fluidized bed and is preferably from about 20 to about 40 vol.%, Most preferably from about 30 to about 35 vol.%

It should be noted that the jet supplied to the contact channel of the source oxygenate can be used as an additional fluidizing gas. Preferably, oxygenate feed

- 4 009365 increased the height of the fluidized bed.

At the upper boundary of the first part of the channel for bringing into contact with the heterogeneous structure of the stream, the oxygenate stream expands, and individual jets merge into one common stream with a uniform structure. A change in the flow structure occurs at the upper boundary of the fluidized bed of catalyst particles. As already noted above, a contact zone with a uniform flow structure begins at the point of the contact introduction channel in which the flow regime is turbulent. The contact zone with the turbulent nature of the flow is located in the zone of maximum expansion of individual oxygenate jets, in which, however, the flow structure still remains inhomogeneous. A zone with a turbulent flow regime is sometimes called an absorption zone. Flow patterns that reduce absorption zones are typically characterized by an increase in contact between the oxygenate and catalyst particles and a decrease in undesirable side reactions. In one embodiment, the height of the absorption zone in the reactor does not exceed 2 feet, preferably does not exceed 0.5 feet.

In a contact zone with a uniform flow structure, the flow rate is preferably from about 18 to about 30 ft / s. Preferably, the concentration of solid catalyst particles in the reaction zone with a uniform flow structure is lower than in the zone with a non-uniform flow structure. At the exit from the reaction zone with a uniform flow structure, i.e. in the place where the stream of oxygenate and catalyst particles enters the connecting channel, the concentration of solid catalyst particles in the stream should be from about 1 to about 3 vol.%.

In the connecting channel, the flow rate of oxygenate and catalyst particles (as well as olefins formed as a result of the conversion of the initial oxygenate) increases. In preferred embodiments, the flow in the connecting channel is homogeneous, and its structure is similar to that of a two-phase flow with plug flow. The plug flow mode in the connecting channel is provided by a certain reduced gas velocity (PSG). PSG is determined by the average velocity of gases flowing through the reactor. Preferably, the PSG in the connecting channel is at least 9 m / s. In one embodiment, the PSG in the last part of the reaction zone is from about 9 to 25 m / s, preferably from 12 to 20 m / s.

III. Types of reactors.

A. Reactors for the conversion of oxygenates to olefins.

The process according to the invention can be used in a wide variety of reactors, such as fluidized bed reactors, fast fluidized bed reactors, circulating fluidized bed reactors, vertical reactors and other types of reactors. Reactors in which the process of the invention can be used are described, in particular, in U8 4076796 and I8 6287522 (dual vertical reactor), as well as Ό. Kipp and O. Lewep8p1e1 in 'ΤΊιιίάίζαΙίοη Espscpeegsd, publisher PoEp E. Kpedeg RiSchyspd Sotrapu, Iete Wogk, Iete Wogk 1977 (all of these publications are hereby incorporated by reference in their entirety).

Preferred types of reactors and vertical reactors in which the process of the invention can be used are described by E.A. 2esh and Ό.Ε. OITO in Kieg ReasYug. И1 and ^ b ^ ζаΐ ^ оη apb Е1шб Рагбс1е 8у51ет5, publishing house Ret1to1b РиЬКЫпд Согрогабоп, Иете Уогк, 1960, р. 48-59, in I8 6166282 (fast fluidized bed reactor) and in I8 09/564613, filed May 4, 2000 (multi-section vertical reactor), all of these publications are fully incorporated into this description by reference.

In one of the preferred from the point of view of practical implementation of embodiments of the invention, the proposed method is carried out in a fluidized bed or a fast fluidized bed using a reactor system, a catalyst regeneration system and a recovery system for olefin (s).

In this embodiment, the reactor system is preferably a fluidized bed reactor system having a first reaction zone located inside one or more vertical reactors, and a second reaction zone located inside at least one apparatus for separating the catalyst from reaction products, usually consisting of one or several cyclones. In one embodiment of the invention, the olefins according to the invention are produced in a plant with one or more vertical reactors and apparatus for separating the catalyst from the reaction products located in a common reactor vessel. The plant-processed feed, preferably containing one or more oxygenates, possibly with one or more diluents, is fed to one or more vertical reactors with particles of a catalytic molecular sieve or a coked catalytic molecular sieve. In this embodiment, prior to being fed to the reactor, the particles of the catalytic molecular sieve or coked catalytic molecular sieve are wetted with a liquid, preferably water or methanol, and / or mixed with a gas, for example an inert gas, in particular nitrogen.

In one embodiment of the invention, the starting oxygenate in a liquid and / or vapor state is supplied to the reactor system in an amount of from 0.1 to about 99.9 wt.%, For example from about 1 to about 99 wt.%, Usually from about 5 to about 95 wt.%, in terms of the entire mass of liquid or gas supplied to the reactor, taking into account the mass of all diluents contained in them.

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The liquid and vaporous starting oxygenates may have the same composition or may contain, in varying proportions, the same or different separate components with the same or different diluents.

The inventive method can be carried out in a wide temperature range, in particular from about 200 to about 1000 ° C, from about 250 to about 800 ° C, or from about 250 to about 750 ° C, mainly from about 300 to about 650 ° C, usually from about 350 to about 600 ° C., preferably from about 350 to about 550 ° C.

Similarly, the proposed invention the method can be carried out in a wide pressure range, including autogenous pressure. Typically, the absolute partial pressure of the feed stream to the reactor of the starting reagents, excluding any diluents contained therein, is from about 0.1 kPa to about 5 MPa, in particular from about 5 kPa to about 1 MPa, usually from about 20 to about 500 kPa.

The mass ratio between the solid particles of the catalyst and the gas in the reactor for conducting solid-state phase reactions usually lies in the range from about 5: 1 to about 75: 1. In a preferred embodiment, this mass ratio is from about 8: 1 to about 50: 1, more preferably from about 10: 1 to about 40: 1.

In the conversion of feed hydrocarbons, preferably containing one or more oxygenates, the amount of olefin (s) obtained, calculated on the total weight of all hydrocarbons obtained, exceeds 50 wt.%, Usually exceeds 60 wt.%, For example, exceeds 70 wt.%, And preferably exceeds 75 wt.%. In one embodiment, the amount of ethylene and / or propylene obtained by the process of the invention, calculated on the total weight of all hydrocarbons obtained, exceeds 65 wt.%, For example, exceeds 70 or 75 wt.%, And preferably exceeds 78 wt.%. Typically, the amount of ethylene produced, based on the weight of all hydrocarbons produced, exceeds 30 wt.%, For example, exceeds 35 wt.% Or even 40 wt.%. In addition, the amount of propylene obtained by the inventive method, calculated on the weight of all hydrocarbons obtained, exceeds 20 wt.%, For example, exceeds 25 or even 30 wt.%, And preferably exceeds 35 wt.%.

The feed oxygenate fed to the reactor is preferably converted, in whole or in part, in the first reaction zone into gaseous reaction products, which, together with the coked catalyst, enter the apparatus for separating it from the reaction products. In a preferred embodiment, a cyclone (s) is used in the apparatus for separating the catalyst from the reaction products, in which the coked catalyst is separated from the gases taken from the reactor containing one or more olefins. In addition to cyclones, the use of which is preferred, methods based on gravitational effects can also be used in the apparatus for separating the catalyst from gaseous reaction products. Plates, caps, elbows and other similar devices may also be used to separate the catalyst from the gaseous reaction products.

In one embodiment of the invention, the apparatus for separating the catalyst from the reaction products has a desorption zone, usually located in its lower part. In the desorption zone, the coked catalyst is brought into contact with a gas, preferably water vapor, methane, carbon dioxide, carbon monoxide, hydrogen or an inert gas, in particular argon, preferably water vapor, or a mixture thereof, and hydrocarbons adsorbed by it are separated from the catalyst, after which the catalyst served in the regeneration system.

The coked catalyst, taken from the apparatus for separating reaction products from it, is fed into the regeneration system. The regeneration system comprises a regenerator, in which the catalyst under normal conditions for regeneration, which include temperature, pressure and residence time of the catalyst in the regenerator, is brought into contact with the regeneration medium, preferably an oxygen-containing gas.

As non-limiting examples of media suitable for catalyst regeneration, oxygen, O ₃ , 8O ₃ , Ν ₂ Ο, N0, ΝΟ ₂ , Ν ₂ Ο ₅ , air, air diluted with nitrogen or carbon dioxide, oxygen and water (patent I8 6245703), carbon monoxide and / or hydrogen, as well as mixtures thereof. The conditions under which it is preferable to regenerate the catalyst should allow coke to burn out of the coked catalyst, preferably to a residual level of less than 0.5% by weight, based on the total weight of the coked catalytic molecular sieve fed to the regeneration system.

So, for example, the temperature in the regenerator should be from about 200 to about 1500 ° C, in particular from about 300 to about 1000 ° C, for example from about 450 to about 750 ° C, preferably from about 550 to about 700 ° C.

The absolute pressure in the regenerator should be from about 15 to about 500 psi (from about 103 to about 3448 kPa), in particular from about 20 to about 250 psi (from about 138 to about 1724 kPa), for example, from about 25 to about 150 psi (from about 172 to about 1034 kPa), preferably from about 30 to about 60 psi (from about 207 to about 414 kPa).

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The residence time of the catalyst in the regenerator can vary widely and range from about 1 minute to several hours, for example from about 1 to 100 minutes. The amount of oxygen in the regenerating off-gas (i.e., in the gas that leaves the regenerator) is usually from about 0.01 to about 5 mol%, based on the total volume of the gas. The amount of oxygen in the gas used to regenerate the coked catalyst (i.e., fresh or supplied to the regenerator gas) is usually at least about 15 mol%, preferably at least about 20 mol%, more preferably from about 20 to about 30 mol .% in terms of the total volume of gas supplied to the regenerator.

Coke burning at the stage of catalyst regeneration is an exothermic reaction, and therefore, in one embodiment of the invention, the temperature in the catalyst regeneration system is continuously controlled by various methods accepted in the art, for example, by supplying chilled gas to the regenerator in a batch, continuous or semi-continuous mode, or a combination of such methods . The preferred method is to withdraw the regenerated catalyst from its regeneration system and pass it through a special cooler to obtain a cooled regenerated catalyst. In one embodiment, a heat exchanger is used as a catalyst cooler located inside or outside the regeneration system. Other catalyst regeneration systems and the methods underlying their operation are described in patent I8 6290916 (regulation of moisture content), which is incorporated herein by reference in its entirety.

The regenerated catalyst after leaving the regeneration system, preferably from its cooler, is mixed with fresh and / or circulating in a closed loop catalytic molecular sieve and / or source oxygenate and / or fresh gas or liquids and returned to the vertical (e) reactor (s) . In one embodiment of the invention, the regenerated catalyst, after exiting the regeneration system, is directly returned to the vertical reactor (s), preferably after passing through an appropriate cooler. To facilitate the supply of the regenerated catalyst to the reactor system, preferably to one or more vertical reactors, it is possible to use a transport gas, for example, an inert gas, oxygenate vapor, water vapor or another carrier gas supplied to the system in semi-continuous or continuous mode.

By controlling the flow rate of the regenerated catalyst or the cooled regenerated catalyst supplied from the regeneration system to the reactor system, the amount of coke on the particles fed to the reactor (s) of the catalytic molecular sieve can be maintained at an optimum level. Various methods for controlling the flow rate of the catalyst supplied to the reactor system are described, for example, in the work of M1syae1 Bond, Exreptei1a1 Tss1tk | is5. S1tsi1ayid Πυίάίζοά Vsl. under the editorship of Stase, Λνίάηη and Kio ^ Joey, publishing house Blaske, 1997, p. 336-337, which is incorporated herein by reference.

The amount of coke on the catalyst, i.e. its coking level is measured on a catalyst taken from the oxygenate conversion process by determining its carbon content. Typically, the amount of coke on the catalytic molecular sieve after its regeneration is from 0.01 to about 15 wt.%, In particular from about 0.1 to about 10 wt.%, For example from about 0.2 to about 5 wt.%, Mainly from about 0.3 to about 2 wt.%, calculated on the total weight of the molecular sieve.

The gaseous stream taken from the system for separating the catalyst from the reaction products is supplied to the reaction product recovery system. Currently, many systems, methods and recovery mechanisms for separating from the gaseous stream taken from the reactor, olefin (s) and its (their) purification are well known. Typically, such recovery systems consist of one or more separation, distillation, and / or distillation towers, columns, dividers, or aggregates of this kind of sequential elements, reaction systems, for example, designed to produce ethylbenzene (patent I8 5476978) and other derivatives, such as aldehydes, ketones and esters (patent I8 5675041), and other relevant equipment, for example various condensers, heat exchangers, cooling systems or units of successive cooling devices, compressors, separators, drum or tanks, pumps and other related equipment.

As non-limiting examples of such towers, columns, dividers or aggregates of this kind of sequential elements, used separately or in combination with each other, one or more demethanizers, preferably operating at high temperature, deethanizers, depropanizers, washing towers, often called caustic wash towers and / or cooling towers, absorbers, adsorbers, membranes, an ethylene separator (C2), a propylene separator (C3) and a boot separator hen (C4).

Various systems that can be used to recover olefin (s), for example ethylene, propylene and / or butene, are described in patents I8 5960643 (production of ethylene enriched secondary stream), I8 5019143, I8 5452581 and I8 5082481 (membrane separation), I8 5672197 (pressure-sensitive adsorbents), I8 6069288 (hydrogen removal), I8 5904880 (single-stage decomposition of

- 7 0093 369 cuppered methanol for hydrogen and carbon dioxide), ϋδ 5927063 (methanol recovery for use in a gas turbine power plant), ϋδ 6121504 (direct quick cooling of the product), ϋδ 6121503 (production of high-purity olefins without ultra-fine distillation) and ϋδ 6293998 pressure), which are fully incorporated into this description by reference.

Other recovery systems with purification systems, such as purification of olefin (s), are described in the encyclopedia Kpk-OIteg Epsus1oreb1a £ Syet1ca1 Tse1to1odu. 4-eb., Ν. 9, Ιοίιη Hubeu & 8op5. 1996, p. 249271 and 894-899, which is incorporated into this description by reference. Cleaning systems are also described in patents ϋδ 6271428 (purification of a stream containing diolefin) and ϋδ 6293999 (separation of propylene from propane), as well as in application ϋδ 09/689363, filed October 20, 2000 (purification of the stream using a hydration catalyst), which are included in the present description by reference.

In conventional recovery systems, additional products, by-products and / or various impurities are formed or accumulate along with the main products. The main products of oxygenate conversion, light olefins, such as ethylene and propylene, are usually purified before further processing into their derivatives, for example, before polymerization. Therefore, in the most preferred embodiments of the recovery systems, a cleaning system is provided. For example, the light (s) olefin (s) obtained (e), in particular by the conversion of methanol, are passed through a purification system in which by-products or various impurities are removed from it (them) to the lowest possible level.

Non-limiting examples of impurities and by-products contained in light olefins include polar compounds such as water, alcohols, carboxylic acids, ethers, carbon oxides, sulfur compounds, for example hydrogen sulfide, carbonyl sulfides and mercaptans, ammonia and other nitrogen compounds , arsine, phosphine and chlorides. Other impurities or by-products include hydrogen and hydrocarbons such as acetylene, methylacetylene, propadiene, butadiene and butine.

Typically, the conversion of one or more oxygenates to olefin (s) with 2 or 3 carbon atoms also produces a small amount of hydrocarbons, in particular olefin (s), with 4 or more carbon atoms. The amount of C ₄₊ hydrocarbons usually does not exceed 20 wt.%, For example 10 or 5 wt.%, Preferably 2 wt.%, Calculated on the total weight of the total gaseous stream taken from the reactor without regard to the mass of water. Typically, a recovery system has one or more reactors designed to convert C ₄₊ impurity hydrocarbons into useful products.

Some non-limiting examples of such reactors are described in patents ϋδ 5955640 (conversion of a product containing four carbon atoms to butene-1), 47δ 4774375 (oligomerization of isobutane and butene-2 to alkylate gasoline), υδ 6049017 (n-butylene dimerization), ϋδ 4287369 and ϋδ 5763678 (the preparation of carbonyl compounds by carbonylation or hydroformylation of high molecular weight olefins with carbon dioxide and hydrogen), υδ 4542252 (multi-stage adiabatic process), ϋδ 5634354 (recovery of olefins and hydrogen), as well as in the work of Soukup 1. et al., Рgo sec £ ог рд рд рд ш ш С C3, C4 б б б C5 O1ePcc δί ^ eat8, Pe1. & Coa1, T. 37, No. 4, 1995 (dimerization or oligomerization of propylene, butylene and pentylene), which are fully incorporated into this description by reference.

B. Particulate matter.

When producing olefins by the method of the invention, the solid particles passing through the reactor for conducting solid-state gas phase reactions in a gas stream are particles of compounds that are solid under normal operating conditions. Such solid particles may be particles of inert materials, catalytic materials, or mixtures thereof. In a preferred embodiment, the solid particles are catalyst particles.

In one embodiment of the invention, the solid particles of the catalyst passing through the inventive reactor system for conducting solid-state gas phase reactions in a gas stream are particles having the properties of a molecular sieve catalyst, in particular a conventional molecular sieve. In this case, zeolitic and non-zeolitic molecular sieves of the large, medium, or narrow pore type can be mentioned as examples. As non-limiting examples of such molecular sieves, narrow-porous molecular sieves of the type AE1, LBT, APC, ΑΤΝ, ATT, ATU, AHUHU, B1K, CΑδ, CHA, CH1, BAC, ΌΌΒ, ΕΌΙ, EK1, LLC, KB1, LIE, BOU, LTA, ΜΟΝ, RAi, PH1, KNO, COO, THO and their substituted forms, medium-pore molecular sieves of the type ABO, AE, EuO, NEi, REC., ME, MB1, MTU, MTT, ΤΟN and their substituted forms , large-porous molecular sieves of the EMT, BAi type and their substituted forms.

Other molecular sieves include the molecular sieves ΑN С, BEA, SB1, CbO, EOH ΟΙδ, Lb, MEK, MOC, MCCA and δΟ ^. As non-limiting the scope of the invention, preferred, in particular for the conversion of oxygenates to olefins, molecular sieves are molecular sieves of the type AE, ABU, BEA, CHA, E1) 1, BAi, BEK, ΟΙδ, LTA, LIT, MEK, MB1, MOC , MTT, MUU, TAM and CURRENT

- 8 009365

In one preferred embodiment of the invention, molecular sieves with a topology of type AE1 or CHA are used, or combinations thereof, preferably molecular sieves with a topology of type CHA.

The crystal structure of all molecular sieves is formed by tetrahedral TO ₄ fragments united by common vertices into a three-dimensional four-bound framework, where T is any tetrahedrally coordinated cation. Such molecular sieves are usually characterized by the size of the pore-bounding ring, which depends on the number of T atoms in the ring. Other characteristics that describe the type of skeleton of a molecular sieve include the location of the rings forming a cavity in the structure of the molecular sieve, as well as the size of the channels in the structure of the molecular sieve, if any, and the gaps between the cavities (see ναη Vekkit et al. RgasIse, 2nd completely revised and supplemented edition, vol. 137, publishing house Eezeg 8s1epse, V.U., Atyegbat, Ie1eg1apb.5. 2001, p. 1-67).

Narrow, medium, and large pore molecular sieves have a framework with 4-12 or more rings. In a preferred embodiment, the molecular sieves have 8-, 10-, or 12-ring structures or structures with a large number of rings and an average pore size of about 3 to 15 A. In the most preferred embodiment, molecular sieves, preferably silicoaluminophosphate molecular sieves, with 8th rings and an average pore size of less than 5 A, preferably from 3 to 5 A, more preferably from 3 to 4.5 A, most preferably from 3.5 to 4.2 A.

Molecular sieves, particularly zeolitic molecular sieves and molecular sieve zeolite-type preferably have a molecular framework formed by one, preferably two or more tetrahedral fragments [TO _4] with common vertices, more preferably two or more tetrahedral fragments [occupies 8 _4], [ A1O ₄ ] and / or [PO ₄ ], most preferably tetrahedral fragments [8X ₄ ], [A1O ₄ ] and / or [PO ₄ ]. Such molecular sieves based on silicon, aluminum and phosphorus and metal-containing molecular sieves based on silicon, aluminum and phosphorus are described in various publications, including, for example,

I8 4567029 (MeARO, where Me means Md, Μη, Ζη or Co),

I8 4440871 (8ARO),

EP 0159624 (EAPAR, where E represents Ak, Be, B, Cr, Co, Ca, Ce, Her, S, Md, Μη, Τι or Ζη),

I8 4554143 (HerAARO),

I8 4822478,

I8 4683217,

I8 4744885 (EYEAR8O),

EP 0158975 and I8 4935216 SCZAR8O),

EP 0161489 (CoAR8O),

EP 0158976 (EAPO, where E represents Co, Her, Md, Μη, Τι or Ζη),

I8 4310440 (A1RO ₄ ),

EP 0158350 (8EYAR8O),

I8 4973460 (LAR8O),

I8 4789535 (IARO),

I8 4992250 (CeAR8O),

I8 4888167 (CeARO),

I8 5057295 (VAR8O),

I8 4738837 (СГАР8О),

I8 4759919 and I8 4851106 (SgARO),

I8 4758419,

I8 4882038,

I8 5434326 and I8 5478787 (MdAR8O),

I8 4554143 (HerAARO),

I8 4894213 (AkAR8O),

I8 4913888 (AkARO),

I8 4686092,

I8 4846956 and I8 4793833 (MpAR8O),

I8 5345011 and I8 6156931 (MpARO),

I8 4737353 (BeAR8O),

I8 4940570 (BeARO),

I8 4801309,

I8 4684617 and I8 4880520 (T1AR8O),

I8 4500651,

I8 4551236 and I8 4605492 (TARO),

I8 4824554 and I8 4744970 (SoAP8O),

I8 4735806 (CaAR8O),

EP 0293937 (OAP8O, where O is the scaffold-forming oxide fragment of YuO ₂ ]),

- 9 0093 365 and υδ 4567029, υδ 4686093, υδ 4781814, υδ 4793984, υδ 4801364, υδ 4853197, υδ 4917876, υδ 4952384, υδ 4956164, υδ 4956165, υδ 497249 530, 50766, 50766 The scope is incorporated into this description by reference.

Other molecular sieves include molecular sieves described in publications

EP 0888187 B1 (microporous crystalline metal phosphates, 8LRO ₄ (UI-6)), υδ 6004898 (molecular sieves and alkaline earth metal), application υδ 09/511943, filed February 24, 2000 (integrated hydrocarbon cocatalyst), application \ ¥ About 01/64340, published September 7, 2001 (thorium-containing molecular sieve),

K. 8/051 as in Nalbok Oö Mo1esi1ag 81ueu8, publ. Wal Yoigapb Keshyob, No. \ ν Wogk, Yeuu Wogk, 1992, which are hereby incorporated by reference in their entirety.

The most preferred molecular sieves containing silicon, aluminum and / or phosphorus and containing aluminum, phosphorus and optionally silicon molecular sieves include aluminophosphate (ALO) and silicoaluminophosphate (8APO) molecular sieves and substituted, preferably metal, molecular sieves ALO and 8APO. Particularly preferred molecular sieves include 8APO molecular sieves and metal substituted 8APO molecular sieves. In one embodiment of the invention, it is proposed to use metal substituted molecular sieves with an alkali metal from group 1A of the periodic table of elements, an alkaline earth metal from group 11A of the periodic table of elements, a rare earth element from group ΙΙΙΒ of the periodic table of elements, including lanthanum, cerium, praseodymium, neodymium , samarium, europium, gadolinium, terbium, dysprosium, holmium, erbium, thulium, ytterbium and lutetium, as well as scandium or yttrium, a transition metal from groups 1УВ, УВ, У1В, У11В, УШВ and ΙΒ Period table of elements or mixtures of any of these metals. In one of the preferred embodiments of the invention, the metal is selected from the group consisting of Co, Cr, Cu, Her, Ca, Ce, Md, Mn, N1, δη, Τι, Ζη and Ζγ, as well as mixtures thereof. In another preferred embodiment, the atoms of the above metals can be embedded in the molecular sieve framework as part of a tetrahedral fragment, such as [MeO ₂ ], and carry a resulting charge, which depends on the valency of the substituent metal. So, for example, in one embodiment of the invention, when the valency of the substituent metal is +2, +3, +4, +5 or +6, the resulting charge of the tetrahedral fragment lies in the range from -2 to +2.

In one embodiment of the invention, it is proposed to use a molecular sieve, which, as described in many of the aforementioned US patents, has the following empirical formula on an anhydrous basis of shK: (M _x A1 _at P,) O2, where K denotes at least one matrix an agent, preferably organic, t means the number of moles of K per 1 mol (M _X A1 _U P ₂ ) O ₂ , the value of which is from 0 to 1, preferably from 0 to 0.5, most preferably from 0 to 0.3, x , y, ζ denote the molar fractions of A1, P and M as tetrahedral oxides, where M is a metal selected from metals of one of the groups 1A, 11A, ΙΒ, ΙΙΙΒ, 1УВ, УВ, У1В, У11В, УСВ and the lanthanide group of the periodic table of elements, preferably from metals of one of the groups including Co, Cr, Cu, Her, Ca, Ce, Md, Mn, N1, δη, Τι, Ζη and Ζγ.

In one embodiment of the invention, it is proposed to use a molecular sieve in which m is greater than or equal to 0.2, and x, Y and ζ are greater than or equal to 0.01.

In another embodiment of the invention, it is proposed to use a molecular sieve, in which m lies in the range from more than 0.1 to 1, x lies in the range from more than 0 to 0.25, y lies in the range from 0.4 to 0.5, and ζ lies in the range from 0.25 to 0.5, more preferably m lies in the range from 0.15 to 0.7, x lies in the range from 0.01 to 0.2, y lies in the range from 0.4 to 0.5, and ζ lies in the range from 0.3 to 0.5.

As non-limiting examples of molecular sieves 8APO and ALPO used in the preparation of olefins by the method of the present invention, mention may be made of one or a combination of molecular sieves of the type 8APO-5, 8APO-8, 8APO-11, 8APO-16, 8APO-17 , 8ARO-18, 8ARO-20, δ ARO-31, 8ARO-34, 8ARO-35, 8ARO-36, 8ARO-37, 8ARO-40, δ ARO-41, 8ARO-42, 8ARO-44 (patent υδ 6162415 ), 8ARO-47, 8ARO-56, ALO-5, ALO-11, ALO-18, ALO-31, ALO-34, ALO-36, ALO-37, ALO-46 and their metal-containing forms. As zeolite molecular sieves, it is preferable to use one or a combination of molecular sieves of the type 8APO-18, δ APO-34, δ APO-3 5, 8APO-44, δAPO-56, ALO-18 and ALO-34, more preferably one or a combination of molecular sieves of the type δAPO-18, δAPO-34, ALO-34 and ALO-18 and their metal-containing forms, most preferably one or a combination of molecular sieves of the type δAPO-34 and ALO-18 and their metal-containing forms.

In one embodiment of the invention, it is proposed to use an intergrowth material, which consists of two or more distinct phases of a crystalline structure, as a molecular sieve. Such fused molecular sieves are described in the application υδ 09/924016, filed August 2001, and in UO 98/15496, published April 16, 1998, which are fully incorporated into this description by reference. Suggested for use in another embodiment.

- 10 009365 implementation of the invention, the molecular sieve has at least one phase of intergrown material with a frame of type ΑΕΙ and CHA. So, for example, the molecular sieves 8 18-18, ALO-18, and K.iD-18 have a типа type skeleton, and the 8ARO-34 molecular sieves have a СНА type skeleton.

In one embodiment of the invention, it is proposed to combine the above molecular sieves with one or more other molecular sieves. In another embodiment, it is proposed to combine silicoaluminophosphate or aluminophosphate molecular sieves or a combination thereof with one or more of the following molecular sieves as non-limiting examples of the invention: Be1a (patent I8 3308069), Ζ8Μ-5 (I8 3702886, ϋδ 4797267 and υδ 5783321), ΖδΜ-11 (υδ 3709979), ΖδΜ-12 (υδ 3832449), ΖδΜ-12 and ΖδΜ-38 (υδ 3948758), ΖδΜ-22 (υδ 5336478), ΖδΜ-23 (υδ 4076842), ΖδΜ-34 (υδΜ-34 (υδΜ-34 (υδΜ-34 (υδΜ-34 (υδΜ-34 ), ΖδΜ-35 (υδ 4016245), ΖδΜ-48 (υδ 4397827), ΖδΜ-58 (υδ 4698217), ΜСΜ-1 (υδ 4639358), ΜСΜ-2 (υδ 4673559), ΜСΜ-3 (υδ 4632811) ΜСΜ-4 (υδ 4664897), ΜСΜ-5 (υδ 4639357), ΜСΜ-9 (υδ 4880611), ΜСΜ-10 (υδ 4623527), ΜСΜ-14 (υδ 4619818), ΜСΜ-22 (υδ 495 4325), ΜСΜ-41 (υδ 5098684), Μ-41δ (υδ 5102643), ΜСΜ-48 (υδ 5198203), ΜСΜ-49 (υδ 5236575), ΜСΜ-56 (υδ 5362697), ALPO-11 (υδ 43104 , titanium aluminum silicates (ΤΑδΟ) of the типаδΟ-45 type (ΕΡ 0229295), borosilicates (υδ 4254297), titanium aluminum phosphates (ΤΑΡΟ) (υδ 4500651), mixtures of ΖδΜ-5 and ΖδΜ-11 (υδ 4229424), ESK-1883 (υδ 4229445) δΑΡΟ-34-linked ALPO-5 (υδ 5972203), a molecular sieve used in the Fischer-Tropsch process (DO 98/57743 dated December 23, 1988), ΜΡΊ-linked zeolites (υδ 6300535) and mesoporous molecular sieves (υδ 6284696 , υδ 5098684, υδ 5102643 and υδ 5108725) (all of these publications are fully incorporated into this description by reference).

Molecular sieve catalysts are prepared or prepared from synthesized molecular sieves and a binder and / or matrix forming material to form a molecular sieve catalyst composition or a finished molecular sieve catalyst composition. Particles of a certain shape and size are then obtained from such a finished molecular sieve-containing catalyst composition by conventional technology, for example by spray drying, granulation, extrusion or other well-known methods.

Various binders can be used to prepare the molecular sieve catalyst composition. As non-limiting examples of binders that can be used individually or in combination with each other, various types of hydrated alumina, silica and / or sol of another inorganic oxide can be mentioned. One of the preferred alumina sols is aluminum chlorohydrol. The inorganic oxide sol has glue properties and binds, mainly during heat treatment, the synthesized molecular sieves with other materials, such as the matrix-forming material. When heated, an inorganic oxide sol, preferably having a low viscosity, is converted into an inorganic oxide matrix. So, for example, an alumina sol turns into an alumina matrix when heated.

Aluminum chlorohydrol, which is a hydroxylated aluminum sol containing chloride as a counterion, has the general formula

A1 _t О _p (ОН) _о С1р-х (Н ₂ О), in which t denotes a number from 1 to 20, η denotes a number from 1 to 8, o denotes a number from 5 to 40, p denotes a number from 2 to 15 , x denotes a number from 0 to 30.

A binder used in one embodiment of the invention, which has the formula A1 ₁₃ O4 (OI) 24C1 ₇ -12 (I2O), is described in Ο.Μ. DoNgtap e1 a1., Δΐιιά. δι.ιιΤ. δει. apy Ca1a1 .. 76, 1993, p. 105-144, which is incorporated herein by reference.

In another embodiment, one or more binders are mixed with one or more alumina materials, examples of which are non-limiting examples of aluminum oxyhydroxide, γ-alumina, boehmite, diaspore and transition aluminas such as α-alumina, β- alumina, γ-alumina, δ-alumina, ε-alumina, κ-alumina and ρ-alumina, aluminum trihydroxides such as gibbside, bairite, nordstrandit, doyelite and mixtures thereof.

In another embodiment of the invention, alumina sols, preferably containing alumina and possibly some silicon, are proposed as binders. In another embodiment, peptized alumina obtained by treating alumina hydrates such as pseudoboehmite with an acid, preferably a halogen-free acid to form sols or solutions of aluminum ions is proposed as binders. As not limiting the scope of the invention, examples of commercially available colloidal alumina sols include Sol No. 11eo 8676 manufactured by the company

- 11 0093 36 No. 11so Siesche1so1 Co., Naperville, units. Illinois, and Sol Juaso1, manufactured by Thier RF Sotrotabop, Valley Forge, pc. Pennsylvania.

In a preferred embodiment, the molecular sieve is used in conjunction with one or more matrix forming materials. The materials forming the matrix usually reduce the total cost of the catalyst and act as a heat sink protecting the catalyst from excessive heat, for example, during its regeneration, increase the density of the catalyst, increase its strength, in particular, increase its crush strength and abrasion resistance, and also control the speed conversions in a specific process.

As non-limiting examples of the matrix forming materials, one or more materials from the group of rare earths, metal oxides, including titanium oxide, zirconium dioxide, magnesium oxide, thorium oxide, beryllium oxide, quartz, silicon dioxide or sols can be mentioned, and mixtures thereof, for example mixtures of silica and magnesium oxide, silica and zirconia, silica and titania, silica and alumina, as well as silica, alumina and thorium oxide. In one embodiment, natural clays, for example clays from the montmorillonite and kaolin families, can also be used as the matrix forming material. Such natural clays include, in particular, sabbentonites and kaolins, known as O | X1C clays. MsIateee, Seogd1a and Nopba. Other non-limiting examples of matrix forming materials include haloysite, kaolinite, dikkit, nakrit or anoxite. In one embodiment, the matrix forming material, preferably clay, is modified by well-known methods, for example, calcining and / or acid treatment and / or chemical treatment.

In one preferred embodiment of the invention, it is proposed to use clay or a clay type composition, preferably clay or a low iron or titanium oxide type clay composition, most preferably kaolin, as the matrix forming material. Kaolin is able to form a pumpable slurry with a high solids content, has a small fresh surface area and is easily compacted due to its plate structure. The average particle size of the matrix-forming material, most preferably kaolin, is preferably from about 0.1 to about 0.6 microns with a particle size distribution at which Ό ₉₀ is less than 1 micron.

In another embodiment, the weight ratio between the binder and the matrix forming material used to prepare the molecular sieve catalyst composition is from 0: 1 to 1:15, preferably from 1:15 to 1: 5, more preferably from 1:10 to 1: 4, most preferably from 1: 6 to 1: 5. When creating the invention, it was found that an increase in the content of the molecular sieve and a decrease in the content of the matrix-forming material increase the catalytic activity of the molecular sieve catalyst composition, while a decrease in the content of the molecular sieve and an increase in the content of the matrix-forming material increase the abrasion resistance of the molecular sieve of the catalyst composition.

In another embodiment of the invention, it is proposed to use a molecular sieve containing catalytic composition in which the molecular sieve is from about 1 to about 99 wt.%, More preferably from about 5 to about 90 wt.%, Most preferably from about 10 to about 80 wt. .% in terms of the total weight of such a composition.

In a further embodiment of the invention, it is proposed to use a spray-dried molecular sieve-containing catalytic composition in or on which the percentage by weight of the binder, calculated on the total weight of the binder, molecular sieve and matrix forming material, is from about 2 to about 30 wt.%, Preferably about from 5 to about 29 wt.%, more preferably from about 7 to about 15 wt.%.

In those cases where the molecular sieve containing the catalytic composition is prepared essentially in dry or dried form, it is usually subjected to heat treatment, such as calcination, at an elevated temperature to further increase its hardness and / or activity. Typically, calcination is carried out in an atmosphere of air with a low water vapor content. The calcination temperature usually ranges from about 400 to about 1000 ° C, preferably from about 500 to about 800 ° C, most preferably from about 550 to about 700 ° C, when it is carried out preferably in an atmosphere of air, nitrogen, helium, flue gases (low oxygen combustion products) or mixtures thereof.

Particles of maleic anhydride-forming catalyst can also be used as a catalyst in the preparation of olefins by the method of the invention. In one embodiment, the maleic anhydride-forming catalyst comprises a mixed oxide with vanadium and phosphorus as its main elements (mixed vanadium-phosphorus oxide). Examples of such catalysts are catalysts containing, as their active component, divadyl pyrophosphate ((UO) ₂ P ₂ O ₇ ). Detailed information on such catalysts and methods for their preparation can be found, for example, in Cb. Keu., 88, 1988, p. 55-80, in DR-B-4-24104 and in I8 4472527 and I8 4520127, which are included

- 12 009365 to the present description by reference.

In the conversion of oxygenates to olefins by the method of the invention, solid inert particles can be present in a stream passing through the reactor system. These inert solid particles can be mixed with the solid particles of the catalyst or used as a carrier for any conventional catalyst. Examples of such inert solid materials are oxides such as silica, alumina, titanium oxide and zirconia. Among these oxides, silicon dioxide is preferred. Silicon dioxide, which, in contrast to other solid materials, is inert in nature, is an extremely effective binder of all components of the oxide catalyst composition and does not reduce the selectivity of many catalyst compositions. In addition, such a material gives the finished catalyst composition a high abrasion resistance. Silica, when used as a catalyst support, can comprise from about 30 to about 70% by weight, preferably from 40 to about 60% by weight, based on the total weight of the catalyst composition and the support.

B. Types of starting materials

In one embodiment of the invention, the chemical reagent is oxygenate, and the catalyst particles are molecular sieves. The use of molecular sieves for the conversion of oxygenate to olefins is preferred. For this purpose, any of the above molecular sieves may be used. It is more preferable to use conventional zeolites and silicoaluminophosphates for the conversion of oxygenates to olefins according to the invention.

Convertible to olefins according to the invention, the oxygenates consist of one or more organic compounds containing at least one oxygen atom. In the most preferred embodiment of the method of the invention, the oxygenate contained in the feedstock is one or more alcohols, preferably aliphatic alcohols, in the aliphatic fragment of which from 1 to 20 carbon atoms, preferably from 1 to 10 carbon atoms, most preferably from 1 to 4 carbon atoms. Alcohols used as feedstock for the production of olefins by the method of the invention include straight and branched chain lower alcohols and corresponding unsaturated alcohols. Examples of oxygenates not limiting the scope of the invention include methanol, ethanol, n-propanol, isopropanol, methyl ethyl ether, dimethyl ether, diethyl ether, diisopropyl ether, formaldehyde, dimethyl carbonate, dimethyl ketone, acetic acid, and mixtures thereof. In a most preferred embodiment, the feedstock is one or more oxygenates from the group consisting of methanol, ethanol, dimethyl ether, diethyl ether and combinations thereof, more preferably from the group comprising methanol and dimethyl ether, and the most preferred starting oxygenate is methanol.

The feedstock in one embodiment of the invention contains one or more diluents, usually designed to reduce the concentration of the feedstock and not react with the feedstock and the catalytic molecular sieve. As non-limiting examples of such diluents, mention may be made of helium, argon, nitrogen, carbon monoxide, carbon dioxide, water, substantially non-reactive paraffins (especially alkanes, such as methane, ethane and propane), substantially non-reactive aromatic compounds and mixtures thereof. The most preferred diluents include water and nitrogen, of which water is a particularly preferred diluent.

The diluent is either directly added to the feedstock at the inlet to the reactor, either fed directly to the reactor, or added together with the molecular sieve catalyst composition. In one embodiment, the amount of diluent in the feed is from about 1 to about 99 mol%, preferably from about 1 to about 80 mol%, more preferably from about 5 to about 50 mol%, most preferably from about 5 up to about 25 mol.% in terms of the total number of moles of feedstock and diluent. In another embodiment of the invention, it is proposed to add to the feedstock either directly or indirectly other hydrocarbons, including olefin (s), paraffin (s), aromatic compound (s) (s) (see, for example, I8 4677242, addition of aromatic compounds) or mixtures thereof, preferably propylene, butylene, pentylene and other hydrocarbons with 4 or more carbon atoms or mixtures thereof.

In yet another embodiment of the invention, the conversion products are used as starting materials processed in a subsequent process. For example, many olefins obtained by the conversion of oxygenates can be used as starting material for the synthesis of polyolefins, such as polyethylene or polypropylene.

IV. Examples.

A. Example installation for the conversion of oxygenates to olefins.

In FIG. 1 shows the lower part 10 of a reactor for producing olefins according to the invention. Shown in FIG. 1, the lower part of the reactor 10 is the lower part of a large vertical reactor of the type described above. At the bottom of the reactor 10

- 13 009365, the inlet channel 100, the contact channel 200, and the connecting channel 300 are located. The catalyst is fed into the inlet channel 100 through a pipe 110. Solid catalyst particles are fed into the inlet channel from the closed loop of the reactor after at least partial regeneration of the catalyst in the described higher regenerator. The input channel 100 is connected to the contact channel 200. The contacting channel 200 consists of a lower truncated cone 210, a middle cylinder 220 and an upper truncated cone 230. In the embodiment shown, the lower truncated cone 210 has an inclined wall 215, and its diameter gradually increases with distance from the pipe 110 through which the reactor feeds the catalyst. The angle 216 of the inclination of the side wall 215 of the lower truncated cone to the horizontal plane is in this case 75 °.

On the side wall 215 there are nozzles 250 through which oxygenates convertible in the reactor are introduced into the contact channel 200. Inside the lower truncated cone 210, the nozzles 250 are inclined at a certain angle to the vertical. The angle of inclination of the nozzles 250 inside the lower truncated cone 210 is indicated by 251. In the embodiment shown in the drawing, this angle 251 is 30 °.

In the contact channel 200, the lower truncated cone 210 is connected to the middle cylinder 220. In the embodiment shown, the lower truncated cone 210 is connected to the cylinder 220 by a flange connection. The connection of the lower truncated cone 210 with the inlet channel 100 is also made flanged. An easily dismountable flange connection will allow the use of different truncated lower truncated cones in the reactor, thereby changing the geometry of the diffuser section located at the inlet. The height of the middle cylinder 220 of the contact channel must be sufficiently large and equal to at least the full height of the contact channel section with the heterogeneous flow structure and the height of the transition section, in which the flow structure is aligned during reactor operation. The middle cylinder of the contact channel has a vertical side wall 225.

The middle cylinder 220 of the contacting channel is connected to the upper truncated cone 230. The upper part of the contacting channel made in the form of a truncated cone 230 gradually reduces the cross-section of the contacting channel to the dimensions of the cross-section of the riser, or connecting channel 300. In the embodiment shown in the drawing the angle 236 of the inclination of the side wall 235 of the upper truncated cone 230 of the contact channel is 85 °. In this case, the distance from the lower truncated cone 210 to the upper truncated cone 230 corresponds to the height of the contact channel. The catalyst, unreacted oxygenates and reaction products rise upward through the connecting channel 300 and enter the separator connected to it.

In the embodiment shown in FIG. 1, the diameter of the inlet channel 100 is 5 feet. The diameter of the middle cylinder 220 of the contact channel is 8.5 feet, and the diameter of the connecting channel 300 is only 6 feet. Such a change in diameter contributes to the formation of the desired flow regime in the reactor. The relatively large diameter of the middle cylinder 220 of the contact channel allows increasing the transit time of individual oxygenate jets through the fluidized bed of catalyst in the channel of contact with the inhomogeneous flow structure until they merge with each other and form one oxygenate stream in the channel of contact. At the exit from the cylindrical portion 220 of the contact channel, by reducing the diameter of the riser to 6 feet, the reduced gas velocity (PSG) increases. With this change in the diameter of the flow part of the reactor, the PSG reaches a minimum of 9 m / s.

During the conversion of oxygenate (s) to olefin (s), solid catalyst particles 505 enter the reactor 10 through the inlet 100. The catalyst particles 505 rising in the reactor in the contacting channel 200 come into contact with the source oxygenate jets 510. As a result of the contact of the oxygenate jets with the stream of solid particles of the catalyst, at least two zones with different flow structures are formed in the contact channel. At the point where the oxygenate jets 510 enter the contacting channel, the individual jets 510 are separated by gaps 525 filled with catalyst particles. In the drawing, this part of the contacting channel is indicated by 520. As the oxygenate jets move upward through the fluidized bed, the individual oxygenate jets gradually merge and form one common gas stream in the mixing zone 530 with catalyst particles suspended therein. The mixed gas stream with catalyst particles rises up and from the contact channel into the connecting channel 300.

B. Modeling the flow in the reactor.

To study the flow structure and flow pattern of the methanol (oxygenate) stream supplied through several inlet nozzles to the contact channel of a reactor designed for the conversion of hydrocarbons into olefins, a model was developed that describes the mixing process of individual jets of methanol. This model was used to model the reactor, the lower part of which is shown in FIG. 1. In this model, the starting methanol was fed into the lower conical part of the contact channel through separate nozzles located around the reactor circumference. In the reactor model, methanol feed nozzles were located in a plane parallel to the base of the reactor below the upper truncated cone 235 of the contact channel. The nozzles were evenly spaced around the circumference of the reactor. Below the plane of the nozzles, they were tilted to the vertical at an angle of 9. Speed

- 14 0093 365 jet of methanol at the exit of the nozzle was taken equal to approximately 96 ft / s. At the exit from the nozzle, the methanol jet had a small diameter, which gradually increased as the jet velocity decreased. At the exit from the nozzles, methanol jets are surrounded by a dense phase consisting of solid catalyst particles suspended in a gas stream. Due to the difference in their concentrations, methanol jets emerging from the nozzles carry away individual solid catalyst particles from the dense phase surrounding them. The percentage of solid particles carried away by methanol streams from the fluidized bed of the catalyst is called the capture coefficient. The volume of the solid fraction in the methanol stream as the jets penetrate the vertical reactor gradually increases. The individual jets merging with each other eventually fill the entire cross section of the reactor. Immediately above this point in the reactor, a classical annular flow structure with an axial flow direction is formed. The kinetics of the reaction occurring in the reactor can be described using the basic physical dependences, which allow one to determine the volume of the jets, the capture coefficient, and the volume of the zone in which the jets merge into one common stream. The following equations were used to simulate the mixing process of the jets and to simulate the structure of the stream in the reactor. These equations show that the key parameters of jet mixing are the coefficient of capture by the jets of solid particles of the catalyst, the angle of inclination of the nozzles to the vertical, and the speed of the methanol stream flowing from the nozzle.

1. Physics of the process of mixing jets.

The position of the center of the jet inside the vertical reactor at any point along the length of 1 jet is determined by the coordinates (Ζ ₀ , Κ <;). The position of the center of the jet changes in accordance with the following equations:

<12 _s _ Tsdoya O <n ^_ u, 'ak _s i. _; ya1po centipoise and 'where represents rate _of discharge of methanol, and ₁ is the jet velocity.

The amount of solid particles carried away by the jet from the catalyst bed is determined by the following equation:

where α is the ablation coefficient, ρ_ί and p _a are the jet density and the density of the emulsion phase, respectively.

The mass balance and the momentum of the jet are rounded off by the following equations: 2ψ (φ ₅ -φ) ar k _} and, '

And [ _etK . ^ Ri, and _r ] = o, ^ - [πΚ'ρ υ, υ _ζ ] = 2ππ -ρ, ψΥ, + πΚ? ₈ (ρ, -RD

The height of the formation zone by separate jets of one common flow (ZOOP) can be determined by the following equation:

And the zoo ^_

where R is the radius of the lower _cylinder (cylindrical) part of the reactor, the second _price tr.ch _ac m _s represents the equivalent diameter of the jets after fusion,

V denotes the speed of the emulsion phase.

Mass balance determines the value of the resulting capture coefficient of solid particles of the catalyst in the confluence of the jets.

- 15 009365

2. Equations describing the geometry of the upper and lower parts of the truncated cone of the contact channel:

Y _ I 1§θ / s s D _{L FINCH} | ' 2 _ - _ _g 1 'conic 2 \ ^{± Ч} Ц11Л ¹ СГОКА / 2 \ ^ХЧ CYL ~ ¹⁴ cyl ¹⁴ riser AND ¹⁴ riser />

V = V + V. copy model

The angle θ in this model is equal to the angle of inclination to the horizontal plane (taper) located between the cylindrical part of the contact channel and the riser of the conical element of the reactor.

Other important reactor operating parameters include the degree of conversion, pressure, temperature, the ratio between the catalyst and hydrocarbons, and the degree of coking of the catalyst, which are respectively 95%, 25 psi, 490 ° C, 30: 1, and 5 wt. .%.

B. Characteristics of the flows in the reactor: the flow structure in the contact channel with the non-uniform flow structure.

In FIG. 2-5 show examples of changes in the state of oxygenate jets in a fluidized bed. These examples are obtained for the reactor model shown in FIG. 1. In the simulation, the starting methanol was injected through 6 nozzles with a diameter of 12 located on the side wall of the reactor with a diameter of 8.5 ft. The methanol velocity at the nozzle exit was 96 ft / s. The gas velocity in the fluidized bed (before methanol was introduced into the reactor) was 2.26 ft / s.

In FIG. 2 shows the movement of the central axis of each jet. The middle curve shows the coordinates of the central axis of the jet obtained as a result of modeling, and two dashed lines show the boundaries of the calculated jet trajectory taking into account the calculation errors. The starting point of the center line of the jet is located in the reactor at zero height at a distance of 4.25 feet from the center axis of the reactor. In the FIGS. 2 graphs show the process of movement of the central axis of the jet in a fluidized bed. The methanol stream passes through the fluidized bed to an upper boundary of the fluidized bed at a height of 21.9 feet, at which the jets begin to intersect with each other.

In FIG. 3 shows the process of expansion of the jets as they move through the fluidized bed of the catalyst. The lower axis shows not the height of the stream in the fluidized bed, but the total distance that the stream passes. At first, the radius of the stream is 0.5 ft. As you move through the fluidized bed, the radius of the stream increases and reaches 1.5 feet.

In FIG. 4 shows the speed of a methanol stream as it moves through a fluidized bed. At the first moment, the jet velocity drops sharply, and then gradually decreases to approximately 20 ft / s.

In FIG. 5 shows the change in volumetric solids content in a methanol stream passing through a fluidized bed. As the methanol stream moves through the fluidized bed, the amount of solid particles present in it continuously increases. At the upper boundary of the fluidized bed, the solids content in the methanol stream increases to about 15 vol.%.

These graphs reflect the typical behavior of methanol jets as they move through the fluidized bed. When moving through a fluidized bed, the jets expand, their speed drops, and the amount of solid particles carried away from them by the fluidized bed increases. The exact values of the height of the fluidized bed and the size, speed and volumetric content in the jets of solid particles of the catalyst may vary depending on the initial speed of the stream and the angle of inclination of the nozzle to the vertical. In this case, however, the qualitative state of the flow structure due to the passage of methanol jets with a lower concentration of solid particles through the fluidized bed remains the same.

D. Characteristics of the flows in the reactor: reactivity during gas-solid phase reactions

To simulate reactivity in a reactor for conducting solid-state gas reactions, the model of the jet mixing process was combined with the kinetic models Ό1 and Ό2 of the reaction proceeding in the reactor. In the combined simulation of the mixing process of the jets and the kinetics of the reaction proceeding in the reactor, each jet was first divided into cylindrical elements of finite volume. The height of each cylindrical element was arbitrarily chosen equal to approximately 1 foot, and the diameter of each element was determined by the model of the mixing process of the jets. The number of finite elements in the range from 9 to 24 was proportional to the axial penetration of the jet into the fluidized bed. The diameter of each finite element was varied taking into account the number of jets (or nozzles). Each element was then assigned an appropriate capture coefficient in accordance with the location of the jet. After that, to determine the moment of jet fusion, as well as the volume and solid phase content in the fusion zone, the geometric criteria necessary for modeling were chosen. At all stages of computer simulation, the operating parameters selected for the basic version, such as the temperature of methanol supplied to the reactor, the ratio between the catalyst and hydrocarbon, the hydrodynamics of the upper zone of the reactor, etc., remained constant. The initial parameters, selected for the model of the jet mixing process, are indicated in the table. Modeling was carried out in order to determine the influence of the coefficient α (coefficient of entrainment or capture by the stream of solid particles of the catalyst), which is the angle of the stream at the inlet of the reactor, and the speed of methanol in the nozzle. During the modeling, the degree of methanol conversion in the lower conical part, the height of the acceleration zone, the capture coefficient, and the selectivity of the reaction were selected as dependent variables.

The value of the ablation coefficient α in a gas stream with suspended solids in it is difficult to determine. For a single-phase flow, a value of 0.08 is indicated in the literature. However, studies of a two-phase flow with solid particles suspended in a gas show that this coefficient is at least two times lower.

The values calculated during the simulation are shown in the table. The average value of the capture coefficient of a stream of solid particles of a catalyst refers to the average concentration of solid particles in a methanol stream passing through a fluidized bed of catalyst. As shown in FIG. 5, the concentration of solid particles in the methanol stream increases as the stream passes through the fluidized bed (see above). In the same way, the capture coefficient in the jet fusion zone refers to the concentration of solid particles in the transition zone (jet fusion zone). The degree of conversion at the bottom of the reactor means the total percentage of starting methanol converted to olefins in the acceleration zone and in the transition zone. In order to avoid excessive methanol conversion (loss of reaction selectivity and the amount of ethylene and propylene produced), the degree of conversion at the bottom of the reactor should not exceed about 90%.

The table shows the effect of the nozzle angle at a fixed value of a equal to 0.04. The value of the entrainment coefficient better reflects the capture process of solid particles of the catalyst passing through a fluidized bed with a stream of methanol. The general tendency is that as the axis of the nozzle approaches the horizontal, the height of the acceleration zone and the lower conversion zone become smaller. (The values indicated in the table for the angle of 15 ° are smaller than for the angle of 30 °, since the boundaries formed by the walls of the vertical reactor at steep nozzle angles change the criteria for jet fusion, i.e., the equivalent diameter of the stream after jet fusion cannot be larger than the diameter of the reactor ) At an inclination angle of 75 °, due to an increase in the height of the junction zone of the jets, the conversion in the lower part of the reactor increases sharply. An increase in the height of the confluence zone affects the conversion reaction more than individual jets. The simulation results presented in the table show that, with a minimum size of the confluence zone for the angle of inclination of the jet at the inlet of the reactor equal to 30 ° to the vertical, with an increase in the height of the fluidized bed, the intensity of the reaction proceeding in the lower zone of the reactor gradually increases. Modeling also showed that as the height of the fluidized bed increases, the size of the confluence zone decreases simultaneously.

The table also includes simulation results with different nozzles. The last column shows the simulation results obtained for nozzles with a diameter of 12.8 with an inclination angle of 30 °. The methanol injection rate through these nozzles was chosen to be 129.4 ft / s. At this speed and an increase in the number of nozzles from 6 to 12, the total volume of methanol fed into the reactor always remained constant. The data in the table show that, at a relatively low height of the jet fusion zone (0.4 ft), the height of the fluidized bed was also relatively small due to the high rate of injection of methanol into it. Therefore, the degree of conversion at the bottom of the reactor was only 0.3296. Despite such a low conversion coefficient at the bottom of the reactor, the process of converting oxygenate to olefins using such nozzles can also be considered as one of the variants of the method proposed in the invention. The low value of the conversion coefficient in this case only means that such nozzles reduce the conversion coefficient in the lower part of the reactor.

Nozzle Diameter (Inch) 12.00 12.00 12.00 12.00 12.00 8.00 Number of nozzles 6 6 6 6 6 12 Nozzle angle fifteen thirty 45 60 75 thirty Nozzle Methanol Speed (ft / s) 96.5 96.5 96.5 96.5 96.5 129.4 Ablation coefficient (a) 0.04 0.04 0.04 0.04 0.04 0.04 The average value of the capture coefficient of a jet of solid particles 0,0535 0,0565 0,0409 0,0286 0.0198 0.0468 Jet Junction Height (ft) 0.0225 0 1,2087 0.7726 3,838 0.4028 Capture coefficient in the jet fusion zone 0.1145 0 0.1288 0.1456 0,1607 0.1 Acceleration Zone Height (feet) 20,0 21.9 17.0 14.2 12.3 12.3 The degree of conversion at the bottom of the reactor 0.5384 0.6237 0.4233 0.3933 0.4358 0.3296

For specialists in this field it is obvious that all of the above embodiments of the invention only illustrate the underlying idea and do not exclude the possibility of making various changes and improvements to them in the scope of the claims.

Claims (10)

1. The method of conversion of oxygenates to olefins, which consists in the fact that a) direct the flow of the catalyst forming olefins through the channel for introducing the catalyst into contact with the oxygenate and feed individual streams of oxygenate into this stream at an angle to its direction, resulting in the conversion of oxygenate to olefins, b) mix the individual jets of oxygenate with each other below the place of their introduction into the catalyst stream relative to the direction of this stream within the specified contact channel and c) direct the resulting mixture into a connecting channel, which has a cross-sectional area less than the cross-sectional area of the contacting channel and in which the conversion of oxygenate to olefins continues. 2. The method according to claim 1, wherein the olefin-forming catalyst stream is supplied to said contact introduction channel from an inlet channel, the contact introduction channel having a section that expands in cross section. 3. The method according to claim 1 or 2, in which individual jets of oxygenate are injected into the channel of introduction into contact with a speed of from 20 to 40 m / s. 4. The method according to any one of the preceding paragraphs, in which the concentration of the olefin-forming catalyst in the connecting channel is less than in the contacting channel. 5. The method according to any one of the preceding paragraphs, in which the solids content in the olefin-forming catalyst in the contact channel is from 15 to 65 vol.%. 6. The method according to any one of claims 1 to 4, in which the solids content in the olefin-forming catalyst in the connecting channel is from 3 to 15 vol.%. 7. The method according to any one of claims 1 to 4, in which the solids content in the olefin-forming catalyst in the connecting channel is from 1 to 5 vol.%. 8. The method according to any one of the preceding paragraphs, in which the obtained olefins are separated from the catalyst forming them. 9. The method of claim 8, in which at least one of the obtained olefin contained in olefins separated from the catalyst is treated with an olefin-forming catalyst to produce a polyolefin. 10. The method according to any one of the preceding paragraphs, in which less than 100% of the total oxygenate supplied to it is converted into olefins in the contact channel.